Catalytic cracking of paraffinic naphtha

ABSTRACT

A gallia-alumina or fluorided gallia-alumina catalyst is used for cracking paraffin-containing hydrocarbon distillate feedstocks to produce light olefins and highly aromatic gasoline.

United States Patent Gale Dec. 16, 1975 CATALYTIC CRACKING F PARAFFINIC2,096,769 /1937 Tropsch 2,889,268 6/1959 Dinwiddie et a1 NAPHTHA3,310,597 3/1967 Goble et a1. 260/683 Inventor: Laird Gale, ust n, Tex.3,770,616 10/1973 Kominami et a1. 208/138 Assignee: She Oil p y, HoustonTex 3,772,184 10/1973 Bertolacinl et a1. 208/65 [22] Filed: Oct. 9, 1973Primary Examiner-Delbert E. Gantz [211 App! 404541 Assistant Examiner-G.E. Schmitkons [52] US. Cl. 208/117; 208/122; 208/135; 208/141; 252/442;252/463; 260/673.5;

260/677 R [57] ABSTRACT [51] Int. C1. ..........................Cl0G11/08; B01] 27/12; I

BOlJ 23/08; C07C 11/02 A gallia-alumina or fluorided gallia-aluminacatalyst is Field of Search 208/115, 1 6, 1 used for crackingparaffin-containing hydrocarbon dis- 208/1 260/6735 tillate feedstocksto produce light olefins and highly aromatic gasoline. [56] ReferencesCited UNITED STATES PATENTS a8 (Ilaims, 1 Drawing Figure 1,935,17711/1933 Connally at al 252/456 CATALYST B E it Q 16 6' CATALYST A 0 I II TIME I MINUTES U.S. Patent Dec. 16, 1975 CATALYST B CATALYST A I T/ME,MINUTES BACKGROUND OF THE INVENTION This invention relates to thecatalytic cracking of hydrocarbons to produce products boiling below theboiling range of the hydrocarbons cracked. In particular it relates tothe catalytic pyrolysis of paraffins to produce light gas and aromatics.

A fluorided alumina catalyst has been used to catalytically crackrefinery feedstocks. A process designed to produce normally-gaseousolefins having a high propylene content is described in U.S. Pat. No.3,310,597.

Substantial aromatization activity results in cracking hydrocarbons overalumina and fluorided alumina, but the yields are poor. A catalyticpyrolysis process which has a high conversion of hydrocarbons to lowerboiling products containing a high yield of light olefins and gasolineboiling range aromatics would be of great value in view of theincreasing demand for these products. Accordingly, it is an object ofthis invention to provide a cracking process which utilizes theintrinsic cyclization-aromatization activity of a modified aluminacatalyst to accomplish such an improved product distribution. Inparticular it is an object of this invention to convert paraffinichydrocarbon feedstocks to lower molecular weight aromatics by a processwhich can be described as "dehydrocracking-aromatization" (DCA).

SUMMARY OF THE INVENTION A catalytic cracking process which comprisescontacting a paraffin-containing hydrocarbon distillate feedstockboiling substantially in the range C,,-450C at cracking conditions witha catalyst comprising a major proportion of alumina combined with about2 to 40% wt gallia to obtain a product containing substantial amounts ofnormally gaseous olefins and gasoline boiling range aromatics. Thecatalyst may optionally contain from about 1 to 5% wt fluoride.

DESCRIPTION OF DRAWINGS The FIGURE shows the effect of catalyst age in aprocess for crackiing cumene to alpha-methylstyrene with alumina andgallia-modified alumina catalysts.

DETAILED DESCRIPTION The present invention is concerned with a catalyticcracking process utilizing a catalyst which is particularly selective inproducinglight normally gaseous olefins and gasoline boiling rangearomatics from paraffincontaining distillate hydrocarbon feedstocksboiling substantially in the range C 450C Such a conversion process,which operates, e. g., by converting a paraffinic feed to a lowermolecular weight aromatic product, can be called adehydroeracking-aromatization (DCA) process.

Since preliminary experiments established that pure alumina exhibits aninteresting cyclization-aromatization activity when cracking normalparaffins, the properties of alumina were studied to see how theyaffected this activity. Two properties were considered most important.The intrinsic acidity of the alumina as described by Pines and coworkersin .I. Am. Chem. Soc., 82, 2471 (1960); and its crystallinemodification, i.e.. eta, gamma, chi. etc.

The most acidic aluminas are prepared from hydrolysis of aluminumisopropoxide while the least acidic aluminas are prepared from sodium orpotassium aluminate. Samples of alumina were prepared by each method andit was determined that both aluminas had the eta crystalline structure.These aluminas were used to crack pure n-octane at a temperature of580C, atmospheric pressure and 1.6 WHSV. These tests showed that theselectivity to aromatics was greater with the nonacidic alumina preparedfrom sodium aluminate. Additional n-octane cracking tests with a gammaand chi alumina showed that the eta form has the highest selectivity toaromatics. Accordingly, the eta form is preferred and has been used inthe galliaalumina catalysts of the invention.

The n-octane cracking studies showed that an alumina with a high alkalimetal content (3.3% wt sodium) had virtually nocyclization-aromatization activity. Accordingly, a low-sodium, etaalumina was used as a standard in determining the effect of reactionvariables on cyclization-aromatization activity. It was concluded thateffective catalysts of the invention should have an alkali metal contentof about 0.2% wt or less.

The gallia-alumina catalysts of the invention may be prepared in variousways, e.g., by impregnation of the alumina with gallia or bycoprecipitation of the gallia and alumina. The latter method isillustrated in Example II below. The catalyst surface area should fallin the range of about 50 to about 300 m /g.

The catalysts of the invention contain a major proportion of alumina,preferably eta alumina, and a minor proportion of gallia, e.g., fromabout 1 to about 40% wt gallia, with from about 3 to about 10% wt galliabeing a preferred composition range.

In addition to gallia and alumina the catlysts of the invention maycontain from about 1 to about 5% wt fluoride to enhance acid crackingactivity of the catalyst. A gallia-alumina catalyst of the inventioncontaining 1.5% wt fluoride was found to be particularly effective incracking a highly paraffinic hydrocracker recycle oil.

The operating temperature range for the process is from about 500 toabout 625C. Conversion increases strongly with temperature whileselectivity to aromatics reaches a maximum in the range of 560 to 590C.The proportion of light gases in the product increases rapidly withincreasing temperature but the quality of the gases is lower, i.e., lessolefinic. Furthermore, the deposition of coke on the catalyst increaseswith increasing temperature.

Suitable operating pressures for the process range from atmospheric (0psig) to about 50 psig. Preferably the pressure ranges from O to about15 psig.

Suitable weight hourly space velocities (WHSV) for the process rangefrom about 0.5 to about 6. Preferably the WHSV will be from about 1 to2. The aromatic content of the gasoline fraction increases markedly withdecreasing WHSV, as does the proportion of light gases. In addition thelight gases become increasingly saturated and thus less valuable. Ingeneral it may be said that reaction conditions which favor the highestaromatic content of the gasoline product yield large gas fractions ofrelatively low olefin content.

. Suitable hydrocarbon feedstocks for the process includeparaffin-containing distillates boiling substantially in the rangeC,,-450C. Preferably the range will be from about C -450C. Since theprocess operates by selectively dehydrocracking-aromatization (DCA) ofparaffins it is necessary that the feedstock contain a substantialproportion of paraffins. Preferably the feedstock will contain fromabout 40 to 100% v paraffins. The process is particularly effective inthe DCA of pure paraffins such as n-octane and n-dodecane. However, thecatalysts of the invention are also useful for the selective DCA ofactual refinery feedstocks having paraffin contents falling within thepreferred range.

When processing such refinery feedstocks it is generally preferred thatfluoride be added to the catalyst to enhance its selectivity toaromatics in the gasoline boiling range.

The catalysts of the invention will generally be applied in a fluid bedprocess where frequent catalyst regenerations are required. A fixed-bedprocess can be used where infrequent regenerations are required.

It has been observed that catalyst aging has three general effects onthe process of the invention:

1. conversion level increases slightly with catalyst age;

2. selectivity to aromatics increases initially (-0.5 hour sample vs.0.5-1.5 hour sample) and then declines somewhat; and

3. the product distribution changes significantly with time. As thecatalyst ages the degree of skeletal isomerization of the C -olefinfraction decreases and the C aromatic product distribution shifts towardo-xylene and ethylbenzene.

Because of the decline in DCA activity with time an alumina catalyst wasprocessed for a 5 hour period with n-dodecane at 0.8 WHSV (to promoterapid coke deposition). Total products were collected and analyzed. Thecatalyst was then regenerated by combustion of the coke with air at580C. The catalyst was then used to process n-dodecane for a 3-hourperiod, during which total products were again collected and analyzed. Acomparison of these test results showed that regeneration hadessentially no effect on conversion level and only a modest (5- l 0%)decrease in selectivity to aromatics. This test suggests that thegallia-alumina catalysts of the invention are amenable to regeneration.

The'invention will now be further illustrated by the following examples.

EXAMPLE 1 An eta alumina Catalyst A was prepared as a basis forcomparison by a method similar to that given by Pines and Haag, J. Am.Chem. Soc. 82, 2471 (1960).

' The preparation method was as follows: 123 g sodium aluminate wasdissolved in 3 liter distilled water. Carbon dioxide was bubbled inuntil no more precipitate was formed. The resulting aluminum hydroxidewas collected by filtration and washed repeatedly with reslurrying toremove sodium ions. The product was dried at l 10C for 2 days. Thealuminum hydroxide was converted to eta alumina by calcining in air at550C for 16 hours. The finished alumina catalyst had a surface area ofabout 210 m /g and a sodium content of about 0.17% wt.

EXAMPLE 11 A gallia-alumina Catalyst B of the invention was prepared asfollows: One mole of aluminum chloride was dissolved in 3 liters ofwater. The pH was adjusted to 7.0 with 6N ammonium hydroxide and asolution of 0.045 mole gallium chloride dissolved in 500 ml water wasadded. After mixing, the pH was further adjusted to 9.5 with ammoniumhydroxide. The resulting gel was aged overnight. The gelled catalyst waswashed six times with 0.1 N NH OH and dried for 4 days at C. Theresulting solid was crushed and meshed to the desired size. Finally, thecatalyst was calcined in air for 16 hours at 550C before use. Thefinished gal1iaalumina Catalyst B had a gallia content of 7.64% wt, asurface area of about 200 m /g and a sodium content of less than 0.1%wt.

EXAMPLE lll Catalysts A and B were used in adehydrocrackingaromatization (DCA) process to convert n-paraffinhydrocarbons to aromatics. The feedstock for this example was puren-dodecane. The feedstock was delivered by a syringe pump to anall-glass reaction system. The reactor was a %inch OD 17-inch long Vycortube. which had a catalyst bed volume of 19 cc and which was heated by athree-section Lindberg l-leviduty Type 705 electric furnace.

A typical catalyst charge consisted of 2 g of catalyst (30-45 mesh)dispersed in 10 g ofquartz chips. A preheat section of the tube wasfilled with quartz chips. Liquid reaction products from the process werecondensed in a water cooled condenser and collected in an efficientglass trap in an ice bath, while gaseous reaction products were takenout of the system through a wet test meter. Representative gas sampleswere collected in glass sampling vessels.

Gaseous reaction products were analyzed by mass spectrometry. The liquidproducts were analyzed by gas-liquid chromatography (GLC) using a 14inchOD x 23.5-foot SF-96/Chromosorb W (acid washed, Hexamethyldisiloxanetreated) column held at 30C for 9 minutes followed by programmed heatingfrom 30 to 250C at 2/minute. Total analysis time is about 2 hours. Peakidentification was accomplished by combining retention time datadeveloped from known compounds and GLC and mass spectrometry analyses.For samples requiring resolution of the C3 aromatics a Ainch OD X20-foot Bentone 34/diisodecylphthalate column was used. Total cokeyields were obtained by a combustion technique.

A Fortran V computer program was written to perform the laboriouscalculations required to combine the gas and liquid product analyses andcoke yield into one overall product distribution. Operating conditionsand test results from these comparative DCA processes are shown in TableI.

Table 1 Temperature: 580C Pressure: Atmospheric Table l-continued Table3 Temperature: 580C SRHGO HRO Pressure: Atmos heric WHSV; 46 Gravity API32.3 44.2 Time: 2.0 hr. Bp Range (GLC). 72w Expt. No. l I 2 Start 82C0.1 Catalyst A B 82 160C 0.8 2.6 C-l2 Aromatics 8.5 2.8 19.3 8 9 160199C 1.3 23.8 199C 216C 1.4 28.3 "'1 mole gallium/21) moles aluminum.216 271C 10.6 31.8 Average of two experiments. 271 86.0 13.5 "Normalizedto 100%. 10 Average Molecular Weight 294 I78 'Moles product/101) molesn-dodecane reacted. Composition. 7rw "C ubstituted benzenes naphthaleneParaft'ms 3 l .3 59 "C -substituted benzenes methylnapthalenesNaphthenes 50.3 32 "C,;-Sl.lblilllld benzenes dimethylandethylnaphthalenes. Aromatics 1 9 U.V. Aromatics, mM/100 g Mono- 56.531.3 Table I shows that the incorporation of gallia into alumina(Catalyst B) increases the selectivity to aromat- :55

Total 62.2 32.3

ics. This is accomplished apparently by increasing the dehydrogenationactivity of the catalyst. Increased dehydrogenation activity shouldincrease the contribution to aromatics formation from dehydrogenation totrienes followed by thermal cyclization. Evidence for enhanceddehydrogenation activity is shown by increased yields of C12 aromaticsfrom n-dodecane.

EXAMPLE IV The enhanced dehydrogenation activity of the galliaaluminaCatalyst B was further demonstrated by cracking cumene in a processsimilar to that described in Example 111. Operating conditions and testresults are shown in Table 2 and the FIGURE. The FIGURE demonstratesthat the yield (wt. plotted as GLC area) of the dehydrogenative productfrom cumene, alphamethylstyrene, is much higher from cumene cracking ongallia-alumina Catalyst B compared to alumina Catalyst A. The initialabsolute yield of benzene from the Catalyst B is nearly comparable tothat from pure alumina indicating comparable concentrations of strongacid sites.

EXAMPLE V Catalysts A and B and a commercial zeolite cracking Catalyst C(Davison DZS were used in a DCA process to crack two refineryfeedstocks. The feedstocks used for these experiments were ahydrotreated straight run heavy gas oil (SRHGO) and a second stagehydrocracked recycle oil (l-IRO) with properties as shown in Table 3.

Table 2 A simple yet accurate test procedure using relatively smallquantities of catalyst and feed was developed to compare catalysts. Theapparatus consisted of a fixedbed microcata-lytic system utilizing theall-glass microflow reactor as described in Example 111. Detailedproduct yield structures were obtained by analyzing product placed inthe Vycor glass reactor. The catalyst bed was heated to 580C with Npurge over a 30 min period and held at 580C for 1 hr with flowing N TheN was then replaced by liquid feed at 7.5 g/hr. Liquid product wascollected for a I hr period. Several representative gas samples werecollected and the total volume of gaseous products was measured using awet test meter. Following product collection, the catalyst bed waspurged for 1 hr with N at 580C. Coke analyses were made by contactingthe catalysts with air and trapping the resulting CO in aqueous sodiumhydroxide. A

heated CuO bed insured completed combustion of CO to CO A titrationprocedure described by Pines and Csiscery, J. of Catalysis 1, 313(1962), was used to determine the carbonate concentration in the aqueoussodium hydroxide solution. Typical material balances of 97% or betterwere obtained using these procedures. The gas samples were analyzed bymass spectrometry while the liquid products were analyzed by temperatureFeed: cumene/Helium N1 Temperature: 580C Pressure: Atmospheric WHSV: 5.8Expt. No. 3 4 Catalyst A" B"' Time. hr 0-1 1-2 01 1-2 ProductDistribution. TLP 7rw *b) %w b) 7rw *b) kw *b) Benzene 5.92 25.0 3.53 .35.39 13.9 1.08 4.6 Toluene 0.39 1.4 0.20 .88 0.59 1.3 Ethylbenzene 3.2810.0 2.41 .0 1.96 3.7 0.45 1.4 Styrene 3.36 10.5 1.91 4 2.20 4.2 0.632.0 Cumene 67.67 72.54 44.56 64.99 n-Propylbenzene 2.67 7.3 2.01 .7 2.053.4 a-Methylstyrene" I 1.33 30.9 13.0 .1 36.60 62.2 30.87 86.5trans-B-Methylstyrene 5.41 14.9 4.42 3 6.65 l 1.3 1.98 5.5 Conversion35.3 5.0 29.5 i 4.3 57.3 35.8

"'Includes some cis-B-merhylstyrene. '"Moles product/100 moles cumenereacted. Surface area. 247 sqm/g.

"'Average of two experiments.

"1 mole gallium/20 moles aluminum programmed GLC. The results of thegas, liquid and coke analyses were combined into a single overallproduct yield structure by a Fortran V computer program.

High boiling products and unconverted components in the feed boilingrange could not be determine'd'directly by GLC. These products weredetermined indirectly by adding an internal marker (%wmethylcyclohexane) and then relating each observed peak area to theknown amount of marker. The difference between the sum of the %w forobserved peaks and 100% .is the amount of undetected higher materials.The accuracy of the internal marker technique was checked for sev-. eralproducts by submitting these samples for GLC boiling point analysiswhich is capable of detecting hydrocarbons up to C The quantities ofproduct boiling 271C determined by the GLC boiling point analysis andthe internal marker technique were in good agreement.

The analysis of the products from cracking the hydrocracker recycle oilwere complicated by the fact that the-feed initial boiling point (23%w160-199C) overlapped the aromatic products in part of the heavy gasolinerange. The yields of benzene, toluenes, ethylbenzene and xylenes couldbe determined directly from the GLC analysis of the TLP. The yields ofhigher alkyl aromatics C -C carbon number) were determined fromhighresolution mass spectral analysis.

The results of cracking tests on the hydrotreated SRHGO feedstockcomparing gallia-alumina Catalyst B of the invention with an eta aluminaCatalyst A and a commercial type zeolite Catalyst C, (Davison DZ-5) areshown in Table 4. Zeolite cracking catalysts are noted for their highconversion yields and selectivity to an aromatic gasoline fraction.

Table 4 Feed: Hydrotreated SRHGO Pressure: Atmospheric Table 4-continuedFeed: Hydrotreated SRHGO Pressure: Atmospheric "1 mole gallium/20 molesaluminum 1 hr. reaction time "'2 hr. reaction time Thesedata show thatthere is no significant difference'i'n the yields of heavy gasoline andthe aromatic contents of the heavy gasoline fraction for the eta aluminaCatalyst A and the zeolite Catalyst C at 580C and atmospheric pressure.However, the hydrogen yield is much higher for Catalyst A indicatinghigher overall aromatization activity for the alumina. The bulk of thisadditional aromatization apparently yields polycondensed aromatics whichend up in the feed boiling range or as coke. The alumina promotedcracking (Catalyst A) yields a poor light olefin distribution. Comparedto zeolite catalytic cracking (Catalyst C), alumina cracking producedconsiderably'more C and C product relative to C and C Also theiso/normal ratio of paraffins is much lower for thelight gas gasolinefraction (C /C from alumina cracking.

The gallia-alumina Catalyst B of the invention, which exhibited enhancedcyclication/aromatization activity with pure n-paraffins, was lessactive than the eta aluggg g v A I g g mina Catalyst A with this SRHGOfeedstock. This poor Temp-emurefc v580 560 580 580 actiyity resultedfrom rapid deactivation by excessive WHSV 1.5" .0" coking. Apparently,the intrinsic cyclization/aromatiza- Product tion activity of Catalyst Bdid not result in a higher yield Diisirilputionfiw 2 2 2 6 l 4 0 6 ofgasoline boiling range aromatics because of the low y rogen 7 Methane 59Z8 42 36 paraffin content (31.3%;v) of thlS feedstock. Ethane 5.0 1.84.0 3.0 The results of cracking tests on the second stage hydrocrackerrecycle oil (HRO) comparing a galliapropylene 13 5,0 72 alumina catalystof the invention with an eta alumina, gu a ne g2 and with fluoridedversions of these catalysts, as well as U V enes 6 HC Gas 285 13] 31,8with the same commercial zeolite cracking catalyst (Davison DZ-5) areshown in Table 5.

Table 5 Temperature: 580C Pressure: Atmospheric WHSV: 1.6 Time: 1.0 hr.Expt. NO. 9 10 11 12 13 Catalyst A A+ 'n B+" C Product Distribution. 7cw

Hydrogen 2.1 1.5 3.2 2.8 0.5 Methane 3.3 4.2 3.0 3.2 2.6 Ethane 2.6 1.72.2 2.0 1.8 Ethylene 2.3 2.9 2.1 4.1 2.0 Propane 1.8 3.2 1.6 1.8 2.4Propylene 4.0 10.8 2.9 6.9 8.7 Butane 1.3 4.0 1.4 3.6 5.4 Butylenes 4.210.4 4.1 7.2 9.2 Sum HC GAS 20.5 37.2 17.3 28.8 32.1 Light Gasoline (C/C) 7 2 s s 5.6 7 2 13 5 Heavy Gasoline (CT/200C) 32.4 20.9 36.0 29.626.6

Table -continued Temperature: 580C Pressure: Atmospheric "'7.647( wtgallia Cracking of the l-IRO on a pure eta alumina (Catalyst A) yields33% more gasoline range aromatics than cracking on the zeolite (CatalystC). Cracking of this practical feedstock over the gallia-alumina(Catalyst B) gave further improvement in the yield of gasolinearomatics.

In order to improve the light olefin product distribution from aluminacracking, the acid cracking activity was increased by incorporatingfluoride into both pure alumina and gallia-alumina catalyst. The data inTable 5 show that cracking the recycle oil on fluorided alumina doesgive improvements in the yields of propylene and butylenes; however, theyield of gasoline range aromatics is adversely affected. On the otherhand, the 1.5%w fluorided gallia-alumina catalyst results in substantialincreases in both propylene and butylene yields while the aromaticsyield is only slightly reduced compared to pure gallia-alumina. Theyield of gasoline range aromatics is 64% greater for Catalyst B 1.5% Fthan that of Catalyst C. The gallia-alumina catalyst B 1.5% F does,however, have a higher coke yield than that for Catalyst C, i.e., 10.2%versus 2.9%.

What is claimed is:

l. A catalyst cracking process which comprises contacting a hydrocarbonfeedstock containing 40 to 100%v paraffms and boiling substantially inthe range C -45OC, at cracking conditions, with a catalyst consistingessentially of alumina containing from about 1 to 40% wt. gallia, andrecovering a product having a major portion of hydrocarbons boilingbelow the boiling range of the feedstock and containing substantialamounts of normally gaseous olefins and gasoline boiling rangearomatics.

2. The process of claim 1 wherein the cracking conditions include atemperature of about 550 to 625C, an operating pressure of 0 to about 50psig, and a weight hourly space velocity of about 0.5 to 6.

3. The process of claim 2 wherein the catalyst consists essentially ofan eta alumina containing about 3 to 10% wt. gallia.

4. The process of claim 2 wherein the catalyst has a surface area fromabout 50 to about 300 m /g.

5. The process of claim 4 wherein the catalyst has a sodium content ofabout 0.2% wt or less.

6. A catalytic cracking process which comprises contacting a hydrocarbonfeedstock containing 40 to %v paraffins and boiling substantially in therange C -450C, at cracking conditions, with a catalyst consistingessentially of alumina containing from about 1 to 40% wt. gallia andfrom about 1 to 5% wt fluoride, and recovering a product having a majorportion of hydrocarbons boiling below the boiling range of the feedstockand containing substantial amounts of normally gaseous olefins andgasoline boiling range aromatics.

7. The process of claim 6 wherein the alumina is eta-alumina, and thealumina contains from about 3 to 10% wt. gallia.

8. The process of claim 6 wherein the hydrocarbon feedstock containsfrom about 40 to 100%v paraffms and boils substantially in the range C-450C, and the catalyst has a surface area from about 50 to about 300 m/g and contains less than about 0.2% wt alkali metal. =l

1. A CATALYST CRACKING PROCESS WHICH COMPRISES CONTACTING A HYDROCARBONFEEDSTOCK CONTANING 40 TO 100%V PARAFFINS AND BOILING SUBSTANTIALLY INTHE RANGE C8-450*C, AT CRACKING CONDITIONS, WITH A CATALYST CONSISTINGESSENTIALLY OF ALUMINA CONTAINING FROM ABOUT 1 TO 40% WT. GALLIA, ANDRECOVERING A PRODUCT HAVING A MAJOR PORTION OF HYDROCARBONS BOILINGBELOW THE BOILING RANGE OF THE FEEDSTOCK AND CONTAINING SUBSTANTIALAMOUNTS OF NORMALLY GASEOUS OLEFINS AND GASOLINE BOILING RANGEAROMATICS.
 2. The process of claim 1 wherein the cracking conditionsinclude a temperature of about 550* to 625*C, an operating pressure of 0to about 50 psig, and a weight hourly space velocity of about 0.5 to 6.3. The process of claim 2 wherein the catalyst consists essentially ofan eta alumina containing about 3 to 10% wt. gallia.
 4. The process ofclaim 2 wherein the catalyst has a surface area from about 50 to about300 m2/g.
 5. The process of claim 4 wherein the catalyst has a sodiumcontent of about 0.2% wt or less.
 6. A catalytic cracking process whichcomprises contacting a hydrocarbon feedstock containing 40 to 100%vparaffins and boiling substantially in the range C8-450*C, at crackingconditions, with a catalyst consisting essentially of alumina containingfrom about 1 to 40% wt. gallia and from about 1 to 5% wt fluoride, andrecovering a product having a major portion of hydrocarbons boilingbelow the boiling range of the feedstock and containing substantialamounts of normally gaseous olefins and gasoline boiling rangearomatics.
 7. The process of claim 6 wherein the alumina is eta-alumina,and the alumina contains from about 3 to 10% wt. gallia.
 8. The processof claim 6 wherein the hydrocarbon feedstock contains from about 40 to100%v paraffins and boils substantially in the range C12-450*C, and thecatalyst has a surface area from about 50 to about 300 m2/g and containsless than about 0.2% wt alkali metal.